Process for the preparation of 1,1,2,2-pentafluoropropane

ABSTRACT

The present invention provides a process for preparing 1,1,1,2,2-pentafluoropropane (245cb), the process comprising gas phase catalytic dehydrochlorination of a composition comprising 1,1,1-trifluoro-2,3-dichloropropane (243db) to produce an intermediate composition comprising 3,3,3-trifluoro-2-chloro-prop-1-ene (CF 3 CCI=CH 2 , 1233xf), hydrogen chloride (HCI) and, optionally, air; and gas phase catalytic fluorination with hydrogen fluoride (HF) of the intermediate composition to produce a reactor product composition comprising 245cb, HF, HCI and air; wherein the process is carried out with a co-feed of air.

The invention relates to a process for preparing1,1,1,2,2-pentafluoropropane (HFC-245cb, referred to hereinafter as245cb). In particular, the invention relates to a process for preparing245cb from 1,1,1-trifluoro-2,3-dichloropropane (HCFC-243db, referred tohereinafter as 243db) via 3,3,3-trifluoro-2-chloro-prop-1-ene(HCFO-1233xf, referred to hereinafter as 1233xf).

245cb is a useful compound, not least as an intermediate in thepreparation of 2,3,3,3-tetrafluoropropene (HFO-1234yf, referred tohereinafter as 1234yf). 245cb is mentioned as an intermediate in thepreparation of 1234yf in WO2009/125199. 245cb is also mentioned inpassing in other documents concerned with the preparation of 1234yf,such as WO2008/054781, WO2013/111911 and US2014/010750.

The listing or discussion of a prior-published document in thisspecification should not necessarily be taken as acknowledgment that thedocument is part of the state of the art or is common general knowledge.

There is a need for an efficient and economic manufacturing process forthe preparation of 245cb. The subject invention address this need by theprovision of a process for preparing 245cb, the process comprising gasphase catalytic dehydrochlorination of a composition comprising 243db toproduce an intermediate composition comprising 1233xf, hydrogen chloride(HCl) and, optionally, air; and gas phase catalytic fluorination of theintermediate composition with hydrogen fluoride (HF) to produce areactor product composition comprising 245cb, HF, HCl and air; whereinthe process is carried out with a co-feed of air.

For the avoidance of doubt, the gas phase catalytic dehydrochlorinationcomprises conversion of the 243db by dehydrochlorination to 1.233xf.Likewise, the gas phase catalytic fluorination comprises conversion of1233xf to 245cb by fluorination. The conversion of 1233xf (CF₃CC1=CH₂)to 245cb (CF₃CF₂CH)₃ involves the addition of two fluorine substituentsand one hydrogen substituent to 1233xf. Put another way, this involvesthe addition of HF and the replacement of the chlorine substituent withfluorine substituent. Thus, the term fluorination (with HF) in thecontext of the subject invention can be considered to include combinedfluorination and hydrofluorination reactions.

The above process may be carried out batch-wise or continuously.Preferably, the process is carried out continuously. The term“continuously” as use herein is intended to include semi-continuousoperation of the process wherein the process is temporarily stopped, forexample, to regenerate and/or replace the catalyst used in the catalyticdehydrochlorination of 243db and/or catalytic fluorination of 1233xf.Certain aspects of the invention enable the cycle time between suchcatalyst regeneration and/or replacement to be lengthened therebyimproving the efficiency and economy of the process.

The catalytic dehydrochlorination of 243db and the catalyticfluorination of 1233xf of the process of the invention can be carriedout together in a single reactor.

In a preferred aspect, however, the catalytic dehydrochlorination of243db and the catalytic fluorination of 1233xf are carried out inseparate first and second reactors, respectively. Typically, there areadvantages associated with the use of separate reactors for these tworeactions, including modifying the conditions in each reactor tofacilitate the catalytic dehydrochlorination of 243db and the catalyticfluorination of 1233xf, respectively. For example, a higher pressure canbe used for the catalytic fluorination of 1233xf compared to thecatalytic dehydrochlorination of 243db. Typically, a somewhat highertemperature can be used in the second reactor compared to the firstreactor. One reason for such a difference in reactor temperature is thepreference for higher temperatures in the second reactor to burn off anycatalyst coking. This is explained in more detail later in thisspecification. The use of two reactors also helps differentconcentrations of HF and air to be used in the catalyticdehydrochlorination and fluorination reactions.

Whether a single or two reactors are used, any suitable apparatus may beused. Typically, the apparatus is made from one or more materials thatare resistant to corrosion, e.g. Hastelloy®, Monel® or Inconel.

Regardless of whether one or two reactors is used, a key feature of theinvention is that it is carried out with a co-feed of air. The inventorshave surprisingly found that this prevents and/or retards coking of thecatalyst or catalysts used in the gas phase catalyticdehydrochlorination of a 243db and/or the gas phase catalyticfluorination of 1233xf (particularly the latter reaction) withoutsignificantly impairing conversion and/or selectivity. Put another way,the use of air has been found to significantly reduce the rate ofcatalyst deactivation in the gas phase transformations of the subjectinvention. This has the effect of lengthening cycle time, which in turnhas benefits of process efficiency and economy. The use of an airco-feed is also believed to enable the process of the invention to beconducted at higher temperatures with a given catalyst. Without beingbound by theory, the air co-feed is thought to help burn coke atapproximately the same rate at which it is produced, thereby extendingcycle time. This is believed to be especially advantageous for the gasphase catalytic fluorination of 1233xf, which compound is fouling to thegas phase fluorination catalyst and relatively difficult to convert to245cb compared to the typically more facile gas phase catalyticdehydrochlorination of a 243db.

It is believed that it is the oxygen in the air that is primarilyresponsible for the unexpected effects described in the precedingparagraph. However, there are advantages to using air in the process ofthe invention rather than oxygen or oxygen enriched air. Use of air(e.g. atmospheric air) is both cheaper than using oxygen or oxygenenriched air. It is also safer to handle air compared to oxygen enrichedair or, particularly, oxygen, due to flammability issues. Theconcentration of oxygen in air (about 21 mol %) is also thought to beespecially suitable for use in the process of the invention, in terms ofthe combination of its effectiveness to prevent and/or retard catalystcoking and its ease of handling. For example, the air, in oneembodiment, is compressed and, optionally, dried, prior to feeding tothe process of the invention. This handling/processing is considerablysafer and more straightforward with air as opposed to oxygen enrichedair or, particularly, oxygen.

In a preferred embodiment, the air is supplied from the atmosphere andis dried prior to entering either reactor. The air may be dried by anydrying method known in the art, but is preferably compressed and thenfed into a drying system comprising a desiccant. Suitable desiccantsinclude silica gel, which can dry the air to a dew point of less thanabout −40° C. In one aspect there are two or more desiccant chambers sothat one can be regenerated whilst the other is drying the air.Alternatively/additionally, the air can be cooled to condense the water.

Typically, the amount of air co-fed to the process of the invention isfrom about 0.1 to about 500 mol %, based on the amount organics fedand/or present in the reactor(s). By organics we mean the carbon-basedcompounds present in the process of the invention, particularly 243db,1233xf and 245cb. In one aspect, the amount of air co-fed to the processof the invention in mol %, as described herein, is based on the amount(i) 243db, (ii) 1233xf, or (iii) the combined amount of 243db and1233xf. In a preferred aspect, for example wherein the process of theinvention is carried out in first and second reactors and the air isco-fed to the second reactor only, the amount of air (mol %) is based onthe amount of 1233xf fed to the second reactor.

Preferably, the amount of air co-fed to the process of the invention isfrom about 1 to about 200 mol %, from about 2 to about 100 mol %, fromabout 5 to about 100 mol % or from about 10 to about 100 mol %, based onthe amount of organics. The preferred amounts of air co-fed to theprocess of the invention are believed to be limited as follows. If toolittle air is used, inadequate prevention and/or retardation of cokingof the catalyst or catalysts is achieved. If too much air is used,selectivity for the desired products is adversely affected and/or thelarge quantities of air become more difficult, and therefore expensive,to handle. It is particularly advantageous that the amount of air co-fedto the process is from about 15 to about 95 mol %, preferably from about20 to about 90 mol %, such as from about 25 to about 85 mol %, based onthe amount of organics. These ranges are currently thought to be optimalfrom the perspective of a combination of ease of handling (e.g. thevolume of air to be handled and its effect on the process design) andability to prevent and/or retard catalyst coking without deleteriouslyaffecting the process chemistry.

When the catalytic dehydrochlorination of 243db and the catalyticfluorination of 1233xf are carried out in separate first and secondreactors, respectively, air may be co-fed to the first reactor and/orthe second reactor. In one embodiment, air is co-fed to the firstreactor and the second reactor, more preferably to the second reactoronly.

Thus, the invention provides a process for preparing 245cb, the processcomprising gas phase catalytic dehydrochlorination in a first reactor ofa composition comprising 243db to produce an intermediate compositioncomprising 1233xf, HF, HCl; and gas phase catalytic fluorination with HFin a second reactor of the intermediate composition to produce a reactorproduct composition comprising 245cb, HF, HCl and air; wherein theprocess is carried out with a co-feed of air to the second reactor.

Whether air is co-fed to the first reactor and the second reactor, or tothe second reactor only, the amount of air co fed to the reactor(s) isbroadly in accordance with the ranges as defined hereinbefore.

However, when air is co-fed to both the first and second reactors, theamount of air co-fed to the first reactor preferably is less than theamount, on a molar basis, of air co-fed to the second reactor. This isbecause 1233xf typically is fouling to the gas phase fluorinationcatalyst in the second reactor and higher concentrations of air arethought to be needed to maintain catalyst stability and activity (e.g.by preventing and/or retarding catalyst coking) in the second reactorcompared to the first reactor. Additionally, more forcing conditions maybe employed in the second reactor compared to the first reactor in orderto achieve the desired levels of 1233xf fluorination conversion andselectivity to 245cb. Higher concentrations of air in the second reactorcompared to the first reactor can help maintain catalyst stability andactivity under such forcing conditions.

Typically, the amount of air co-fed to the first reactor is less thanhalf the amount co-fed to the second reactor, preferably less than aquarter of the amount of air co-fed to the second reactor, such as lessthan a tenth of the amount of air co-fed to the second reactor. By wayof example, when air is co-fed to both the first and second reactors,the amount of air co-fed to the first reactor typically is from about0.1 to about 100 mol %, preferably from about 0.2 to about 50 mol %,such as from about 0.3 to about 20 mol %, for example from about 0.4 toabout 10 mol %, based on the amount of organics (e.g. based on 243db);whereas the amount of air co-fed to the second reactor typically is fromabout 1 to about 200 mol %, preferably from about 5 to about 100 mol %,such as from about 10 to about 90 mol %, for example from about 15 toabout 85 mol %, based on the amount of organics (e.g. based on 1233xf).

In a preferred embodiment when the catalytic dehydrochlorination of243db and the catalytic fluorination of 1233xf are carried out inseparate first and second reactors, the intermediate composition exitingthe first reactor is fed directly to the second reactor. This has theadvantage of process economy. For example, when air is co-fed to thefirst reactor, the intermediate composition contains air. It ispreferable, when air is co-fed to the first reactor, for air also to befed to the second reactor. This can be achieved simply by feeding theintermediate composition exiting the first reactor directly to thesecond reactor without an intermediate purification step (e.g. to removeair and/or Hel). Is has been found that the presence of HCl does notsignificantly disadvantage the fluorination of 1233xf. The unexpectedbenefit of this is the intermediate composition exiting the firstreactor can be fed directly to the second reactor without removing HCl,which removal requires energy to cool the composition, remove HCl andre-heat the composition. Of course, even when feeding the intermediatecomposition exiting the first reactor directly to the second reactorwithout an intermediate purification step, it may be desirable to heator cool the intermediate composition, for example if the fluorinationreaction in the second reactor is being carried out at a highertemperature than the dehydrochlorination reaction in the first reactor.In the embodiment wherein the intermediate composition is fed directlyto the second reactor, it is preferable to have an additional co-feed ofair into the second reactor because, as explained above, higherconcentrations of air in the second reactor compared to the firstreactor can help prevent and/or retard catalyst coking.

The catalyst used in the catalytic dehydrochlorination step may be anysuitable catalyst that is effective to dehydrochlorinate 243db.Preferred catalysts are bulk form or supported catalysts comprisingactivated carbon, a zero-valent metal, a metal oxide, a metal oxyhalide,a metal halide, or mixtures of the foregoing.

For the avoidance of doubt, by bulk form or supported catalysts,catalysts comprising activated carbon, a zero-valent metal, a metaloxide, a metal oxyhalide, a metal halide, or mixtures of the foregoing,we include catalysts that are essentially only bulk form or supportedcatalysts, catalysts comprising activated carbon, a zero-valent metal, ametal oxide, a metal oxyhalide, a metal halide, or mixtures thereof, andsuch catalysts that are modified, for example, by the addition of one ormore promoters or excipients. Suitable promoters include metals (e.g.transition metals) and/or compounds thereof, and suitable excipientsinclude binders and/or lubricants.

By “activated carbon”, we include any carbon with a relatively highsurface area such as from about 50 to about 3000 m² or from about 100 toabout 2000 m² (e.g. from about 200 to about 1500 m² or about 300 toabout 1000 m²). The activated carbon may be derived from anycarbonaceous material, such as coal (e.g. charcoal), nutshells (e.g.coconut) and wood. Any form of activated carbon may be used, such aspowdered, granulated and pelleted activated carbon. Activated carbonwhich has been modified (e.g. impregnated) by the addition of Cr, Mn,Au, Fe, Sn, Ta, Ti, Sb, Al, Co, Ni, Mo, Ru, Rh, Pd and/or Pt and/or acompound (e.g. a halide) of one or more of these metals may be used.

Suitable catalysts comprising a zero-valent metal including supported(e.g. by carbon) transition metals such as Pd, Fe, Ni and Co.

Suitable metals for the catalysts comprising a metal oxide, a metaloxyhalide or a metal halide include transition metals, alkaline earthmetals (e.g. Mg) and main group metals such as Al, Sn or Sb.

Alumina which has been modified by the addition of Cr, Cu, Zn, Mn, Au,Fe, Sn, Ta, Ti, Sb, In, Co, Ni, Mo, Ru, Rh, Pd and/or Pt and/or acompound (e.g. a halide) of one or more of these metals may be used.

A further group of preferred catalysts are supported (e.g. on carbon)lewis acid metal halides, including TaX₅, SbX₅, SnX₄, TiX₄, FeCl₃,NbXs₅, VX₅, AlX₃ (wherein X═F or Cl). An oxide of a transition metalthat has been modified by the addition of Cr, Mn, Au, Fe, Sn, Ta, Ti,Sb, In, Al, Co, Ni, Nb, Mo, Ru, Rh, Pd and/or Pt and/or a compound (e.g.a halide) of one or more of these metals may be used.

A preferred oxide of a transition metal is an oxide of Cr, Ti, V, Zr, orFe. For example, chromia (Cr₂0₃) alone or chromia that has been modifiedby the addition of Zn, Mn, Mo, Nb, Zr, In, Ni, Al and/or Mg and/or acompound of one or more of these metals may be used. Catalysts based onchromia currently are particularly preferred. A preferred chromia-basedcatalyst is a zinc/chromia catalyst.

By the term “zinc/chromia catalyst” we mean any catalyst comprisingchromium or a compound of chromium and zinc or a compound of zinc. Suchcatalysts are known in the art, see for example EP-A-0502605,EP-A-0773061, EP-A-0957074, WO 98/10862, WO 2010/116150, which documentsare incorporated herein by reference.

Typically, the chromium or compound of chromium present in thezinc/chromia catalysts of the invention is an oxide, oxyfluoride orfluoride (preferably an oxide or oxyfluoride) of chromium.

The total amount of the zinc or a compound of zinc present in thezinc/chromia catalysts of the invention is typically from about 0.01% toabout 25%, preferably 0.1% to about 25%, conveniently 0.01% to 6% zinc,and in some embodiments preferably 0.5% by weight to about 25% by weightof the catalyst, preferably from about 1 to 10% by weight of thecatalyst, more preferably from about 2 to 8% by weight of the catalyst,for example about 4 to 6% by weight of the catalyst. In otherembodiments, the catalyst conveniently comprises 0.01% to 1%, morepreferably 0.05% to 0.5% zinc. It is to be understood that the amount ofzinc or a compound of zinc quoted herein refers to the amount ofelemental zinc, whether present as elemental zinc or as a compound ofzinc.

The zinc/chromia catalysts used in the present invention may beamorphous. By this we mean that the catalyst does not demonstratesubstantial crystalline characteristics when analysed by, for example,X-ray diffraction. Alternatively, the catalysts may be partiallycrystalline. By this we mean that from 0.1 to 50% by weight of thecatalyst is in the form of one or more crystalline compounds of chromiumand/or one or more crystalline compounds of zinc. If a partiallycrystalline catalyst is used, it preferably contains from 0.2 to 25% byweight, more preferably from 0.3 to 10% by weight, still more preferablyfrom 0.4 to 5% by weight of the catalyst in the form of one or morecrystalline compounds of chromium and/or one or more crystallinecompounds of zinc.

The percentage of crystalline material in the catalysts of the inventioncan be determined by any suitable method known in the art. Suitablemethods include X-ray diffraction (XRD) techniques. When X-raydiffraction is used the amount of crystalline material such as theamount of crystalline chromium oxide can be determined with reference toa known amount of graphite present in the catalyst (e.g. the graphiteused in producing catalyst pellets) or more preferably by comparison ofthe intensity of the XRD patterns of the sample materials with referencematerials prepared from suitable internationally recognised standards,for example NIST (National Institute of Standards and Technology)reference materials.

The zinc/chromia catalysts typically have a surface area of at least 50m²/g and preferably from 70 to 250 m²/g and most preferably from 100 to200 m²/g before it is subjected to pretreatment with a fluoridecontaining species such as hydrogen fluoride or a fluorinatedhydrocarbon. During this pre-treatment, which is described in moredetail hereinafter, at least some of the oxygen atoms in the catalystare replaced by fluorine atoms.

The amorphous zinc/chromia catalysts which may be used in the presentinvention can be obtained by any method known in the art for producingamorphous chromia-based catalysts. Suitable methods includeco-precipitation from solutions of zinc and chromium nitrates on theaddition of ammonium hydroxide. Alternatively, surface impregnation ofthe zinc or a compound thereof onto an amorphous chromia catalyst can beused.

Further methods for preparing the amorphous zinc/chromia catalystsinclude, for example, reduction of a chromium (VI) compound, for examplea chromate, dichromate, in particular ammonium dichromate, to chromium(III), by zinc metal, followed by co-precipitation and washing; ormixing as solids, a chromium (VI) compound and a compound of zinc, forexample zinc acetate or zinc oxalate, and heating the mixture to hightemperature in order to effect reduction of the chromium (VI) compoundto chromium (III) oxide and oxidise the compound of zinc to zinc oxide.

The zinc may be introduced into and/or onto the amorphous chromiacatalyst in the form of a compound, for example a halide, oxyhalide,oxide or hydroxide depending at least to some extent upon the catalystpreparation technique employed. In the case where amorphous catalystpreparation is by impregnation of a chromia, halogenated chromia orchromium oxyhalide, the compound is preferably a water-soluble salt, forexample a halide, nitrate or carbonate, and is employed as an aqueoussolution or slurry. Alternatively, the hydroxides of zinc and chromiummay be co-precipitated (for example by the use of a base such as sodiumhydroxide or ammonium hydroxide) and then converted to the oxides toprepare the amorphous catalyst. Mixing and milling of an insoluble zinccompound with the basic chromia catalyst provides a further method ofpreparing the amorphous catalyst precursor. A method for makingamorphous catalyst based on chromium oxyhalide comprises adding acompound of zinc to hydrated chromium halide.

The amount of zinc or a compound of zinc introduced to the amorphouscatalyst precursor depends upon the preparation method employed. It isbelieved that the working catalyst has a surface containing cations ofzinc located in a chromium-containing lattice, for example chromiumoxide, oxyhalide, or halide lattice. Thus the amount of zinc or acompound of zinc required is generally lower for catalysts made byimpregnation than for catalysts made by other methods such asco-precipitation, which also contain the zinc or a compound of zinc innon-surface locations.

The catalysts described herein (e.g. the chromia-based catalysts such aszinc/chromia catalysts) are typically stabilised by heat treatmentbefore use such that they are stable under the environmental conditionsthat they are exposed to in use. This stabilisation is often a two-stageprocess. In the first stage, the catalyst is calcined by heat treatmentin nitrogen or a nitrogen/air environment. The catalyst is thentypically stabilised to hydrogen fluoride by heat treatment in hydrogenfluoride. This stage is often termed “prefluorination”.

By careful control of the conditions under which these two heattreatment stages are conducted, crystallinity can be induced into thecatalyst to a controlled degree.

In use, the catalysts described herein (e.g. the chromia-based catalystssuch as the zinc/chromia catalysts) may be regenerated or reactivatedperiodically by heating in air at a temperature of from about 300° C. toabout 500° C. Air may be used as a mixture with an inert gas such asnitrogen or with hydrogen fluoride, which emerges hot from the catalysttreatment process and may be used directly in any fluorination processesemploying the reactivated catalyst.

The vapour phase catalytic dehydrochlorination may be carried out at atemperature of from about 200 to about 450° C. and at atmospheric, sub-or super-atmospheric pressure, preferably from about 0.1 to about 30bara. Preferably, the catalytic dehydrochlorination is conducted at atemperature of from about 250 to about 400° C., such as from about 280to about 380° C. or from about 300 to about 350° C.

The vapour phase catalytic dehydrochlorination preferably is carried outat a pressure of from about 0.5 to about 25 bara or about 1 to about 20bara, such as from about 2 to about 18 bara (e.g. about 5 to about 20bara or about 8 to about 18 bara or about 10 to about 15 bara).

HF is required for the fluorination of 1233xf in the process of theinvention. The molar ratio of HF: 1233xf in the catalytic fluorinationstep is typically from about 1:1 to about 45:1, such as from about 1:1to about 30:1, preferably from about 1.5:1 to about 30:1, such as fromabout 2:1 to about 20:1 or from about 3:1 to about 15:1. The inventorshave unexpectedly found that these ranges strike a balance between thedesirability to prevent and/or retard catalyst coking and residencetime. If too little HF is used, coking increases. If too much HF isused, the residence time for a given reactor volume becomes shorter thandesired.

If the catalytic dehydrochlorination and fluorination reactions arecarried out in the same reactor, then both reactions are carried out inthe presence of HF. When first and second reactors are used for thecatalytic dehydrochlorination and fluorination reactions, then thereneed not be any HF present in the first reactor for the catalyticdehydrochlorination reaction. However, in some embodiments it is thoughtpreferable to have HF present for the catalytic dehydrochlorination.Without being bound by theory, this is believed to prevent and/or retardcatalyst coking.

If HF is present in the catalytic dehydrochlorination step, the molarratio of HF:243db can fall within the ranges defined above for the molarratio of HF:1233xf in the catalytic fluorination of 1233xf. In oneaspect, however, less HF is used in the catalytic dehydrochlorinationstep compared to the catalytic fluorination step. Thus, the compositioncomprising 243db can additionally contain HF, typically in a molar ratioof HF:243db of from about 0.5:1 to about 40:1, such as from about 0.5:1to about 20:1, preferably from about 1:1 to about 15:1, such as fromabout 1.5:1 to about 10:1 or from about 2:1 to about 8:1.

The contact time for the composition comprising 243db and HF with thecatalyst in the catalytic dehydrochlorination step typically is fromabout 0.5 to about 200 seconds, such as from about 1 to about 150seconds. Preferably, the contact time is from about 1 to about 100seconds, such as from about 2 to about 80 seconds or from about 8 toabout 60 seconds.

Turning now to the gas phase catalytic fluorination of the intermediatecomposition of the process of the invention, the HF in the intermediatecomposition typically is used to fluorinate 1233xf to 245cb. Preferably,the HF in the intermediate composition is the sole fluorinating agentfor conversion of 1233xf to 245cb, although additional HF can be addedto the process of the invention to facilitate this, particularly if asecond reactor is used for the catalytic fluorination of 1233xf.

The catalyst used in the catalytic fluorination step may be any suitablecatalyst that is effective to fluorinate 1233xf to 245cb. Preferredcatalysts are bulk form or supported catalysts comprising activatedcarbon, a zero-valent metal, a metal oxide, a metal oxyhalide, a metalhalide, or mixtures of the foregoing as described above in relation tothe catalyst for the catalytic dehydrochlorination step.

Preferred catalysts for catalytic fluorination of 1233xf to 245cb arethose which comprise chromia, alone or chromia that has been modified bythe addition of Zn, Mn, Mo, Nb, Zr, In, Ni, Al and/or Mg and/or acompound of one or more of these metals. A preferred chromia-basedcatalyst for use in the catalytic fluorination of 1233xf to 245cb is azinc/chromia catalyst. The same catalyst (e.g. a chromia-based catalyst)may be used for the catalytic dehydrochlorination and fluorinationsteps.

The vapour phase catalytic fluorination step may be carried out at atemperature of from about 200 to about 450° C. and at atmospheric, sub-or super-atmospheric pressure, preferably from about 0.1 to about 30bara. Preferably, the vapour phase catalytic fluorination is conductedat a temperature of from about 250 to about 420° C., such as from about280 to about 400° C. or from about 300 to about 380° C. (e.g. from about330 to about 380° C.).

The vapour phase catalytic fluorination preferably is carried out at apressure of from about 0.5 to about 25 bara or about 1 to about 20 bara,such as from about 2 to about 20 bara (e.g. about 5 to about 20 bara orfrom about 1.0 to about 15 bara).

The contact time for the for the composition comprising 1233xf, HCl andHF with the catalyst in the catalytic fluorination step typically isfrom about 0.5 to about 200 seconds, such as from about 1 to about 150seconds. Preferably, the contact time is from about 1 to about 100seconds, such as from about 2 to about 80 seconds or from about 5 toabout 50 seconds.

245cb is a useful starting material for the manufacture of 1234yf.Accordingly, the process of the invention further comprises feeding245cb into a dehydrofluorination reactor to produce adehydrofluorination product comprising 2,3,3,3-tetrafluoropropene(1234yf) and HF.

The dehydrofluorination of 245cb may be carried out in the vapour and/orliquid phase and typically is carried out at a temperature of from about−70 to about 1000° C. (e.g. 0 to 450° C.). The dehydrofluorination maybe carried out at atmospheric sub- or super atmospheric pressure,preferably from about 0.1 to about 30 bara.

The dehydrofluorination may be induced thermally, may be base-mediatedand/or may be catalysed by any suitable catalyst. Suitable catalystsinclude metal and carbon based catalysts such as those comprisingactivated carbon, main group (e.g. alumina-based catalysts) andtransition metals, such as chromia-based catalysts (e.g. zinc/chromia),lewis acid metal halides or zero-valent metal catalysts. One preferredmethod of effecting the dehydrofluorination of the compound of 245cb toproduce 1234yf is by contacting with a metal-based catalyst, such as achromia-based (e.g. zinc/chromia) catalyst.

Preferably, the 245cb is catalytically dehydrofluorinated to 1234yf inthe gas phase.

243db is commercially available (e.g. from Apollo Scientific Ltd, UK).Alternatively, 243db may also be prepared via a synthetic route startingfrom the cheap feedstocks carbon tetrachloride (CCl₄) and ethylene (seethe reaction scheme set out below). These two starting materials may betelomerised to produce 1,1,1,3-tetrachloropropane (see, for example, J.Am. Chem. Soc. Vol. 70, p 2529, 1948, which is incorporated herein byreference) (also known as HCC-250fb, or simply 250fb).

250fb may then be fluorinated to produce 3,3,3-trifluoropropene (1243zf)and/or 1,1,1-trifluoro-3-chloropropane (253fb) (e.g. using HF,optionally in the presence of a chromiacontaining catalyst, preferably azinc/chromia catalyst as described herein). Dehydrohalogenation of1,1,1-trifluoro-3-chloropropane (e.g. using NaOH or KOH or in the vapourphase) produces 3,3,3-trifluoropropene (1243zf).

1243zf may then be readily halogenated, such as chlorinated (e.g. withchlorine) to produce 1,1,1-trifluoro-2,3-dichloropropane (243db). Thisreaction scheme is summarized below.

The preparation of 243db outlined above is described in more detail inWO 2010/116150 and WO 2009/125199, which are incorporated herein byreference.

Embodiments of the present invention will now be described withreference to the following non-limiting examples and drawings:

FIG. 1 shows a schematic process flow sheet in accordance with theinvention;

FIG. 2 shows the results of a coking study in which conversion isplotted over time for the fluorination of 1233xf in accordance with theinvention.

FIG. 1 illustrates a process design in accordance with the invention. Acomposition (1) comprising 243db and HF is introduced into a firstreactor (A) in which gas phase catalytic dehydrochlorination occurs toproduce an intermediate composition (2) comprising 1233xf, HF and HCl.The intermediate composition may further contain unreacted 243db and, incertain embodiments, by-products such as 245cb and 1234yf.

The intermediate composition (2) is fed directly to a second reactor(B), as is a co-feed (3) of air, and gas phase catalytic fluorination ofthe intermediate composition (2) occurs in the second reactor (B) toproduce a reactor product composition (4) comprising 245cb, HF, HCl andair. The reactor product composition may further contain unreacted1233xf and, in certain embodiments, unreacted 243db and by-products suchas 1234yf.

In a preferred embodiment, the reactor product composition (4) isseparated at separation step (C) into a stream (5) comprising HCl andair and a stream (6) comprising 245cb and HF. An advantage of the use ofthe co-feed of air in the process of the invention is that it can bereadily separated from the reactor product composition together withHCl. Preferably, this is achieved by distillation, with the stream (5)comprising HCl and air being taken off the top of a distillation column(e) and the stream (6) comprising 245cb and HF being taken off thebottom of the distillation column (e). The stream (6) typically containsany other components present, such as unreacted 243db, 1233xf and/or1234yf.

In the embodiment illustrated by FIG. 1, the stream (6) comprising 245cband HF is separated at separation step (D) into a 245cb-rich stream (7)and a HF-rich stream (8). Preferably, this is achieved by distillation,with the 245cb-rich stream (7) being taken off the top of a distillationcolumn (D) and the HF-rich stream (8) being taken off the bottom of thedistillation column (D). The 245cb-rich stream (7) typically alsocontains any relatively light organic components present, such as1234yf. The HF-rich stream (8) typically also contains any relativelyheavy organic components present, such as 1233xf.

Preferably the 245cb-rich stream (7) is subjected to a scrubbing step(E) in which any residual HF (and/or indeed any residual HCl) issubstantially removed from the 245cb-rich stream to produce a 245cb-richstream (11) substantially free from HF (and/or substantially free fromHCl). Typically, this step (E) involves contacting the 245cb-rich stream(7) with water and/or or with a source of aqueous acid and/or or with asource of aqueous alkali, generally represented in FIG. 1 as stream (9),to generate the 245cb-rich stream (11) substantially free from HF andone or more spent scrubbing streams (10). By substantially free from HF,we include the meaning of less than 100 ppm, preferably less than 50ppm, 40 ppm, 30 ppm, 20 ppm, 10 ppm, 5 ppm 4, ppm, 3 ppm, 3 ppm or lessthan 1 ppm. In a preferred embodiment, the 245cb-rich stream (11) issubjected to a separation step (F) in which the 245cb is furtherseparated from any further organic components present (e.g.fluorocarbons such as 1234yf) to produce a substantially pure 245cbproduct (13). Preferably, this separation step (F) comprises one or moredistillation steps. By substantially pure 245cb product (13), we includethe meaning of greater than 95%, 98%, 99% pure, preferably greater than99.5%, 99.8% or 99.9% pure, on a molar basis.

In a preferred embodiment, the HF in the HF-rich stream (8) is recycledto the catalytic dehydrochlorination of the composition comprising 243dband HF. As shown in FIG. 1, preferably, the HF-rich stream is subjectedto a separation step (G) in which the HF-rich stream (8) is separatedinto an HF stream (14) and an organic stream (15). The HF stream isrecycled to the composition (1) comprising 243db and HF which enters thefirst reactor (A) in which gas phase catalytic dehydrochlorinationoccurs. In a preferred aspect, the separation step (G) comprises a phaseseparator.

EXAMPLE 1

A series of catalysts (see Table 1 below) were screened for 243dbdehydrochlorination. The test catalysts were ground to 0.5-1.4 mm and 2mL was charged to an Inconel 625 reactor (0.5″ OD×32 cm). The catalystswere pre-dried at 200° C. for at least 2 hours under a flow of N₂ (60ml/min). All the catalysts shown, except activated carbon, werepre-fluorinated as follows. HF at 30 ml/min was passed over the catalystalong with 60 ml/min nitrogen at 300° C. for one hour. The nitrogen wasdirected to the reactor exit leaving neat HF passing over the catalyst.The temperature was slowly ramped to 360° C. and held for hours beforereducing to 250° C. All the experiments were run at atmospheric pressureand at the temperatures indicated. The 243db flow was 2 ml/min withactivated carbon catalyst and ranged from 0.5 to about 1 ml/min for theremaining catalysts. All experiments were conducted with an HF flow inexcess of the 243db flow, except for the activated carbon catalyst runs,in which no HF was used. Reactor off-gas was sampled scrubbing throughdeionised water and analysed by gas chromatography. The 243db conversionand selectivity to 1233xf are shown in Table 1.

TABLE 1 Experimental results for 243db dehydrochlorination Temperature243db 1233xf Catalyst ° C. Conversion % selectivity % 2% Zn/Chrome 25098.22 74.79 300 100.00 42.28 350 100.00 68.74 4% Zn/Chrome 250 99.1881.80 300 98.47 77.42 350 100.00 70.80 6% Zn/Chrome 250 94.94 65.55 300100.00 33.81 350 100.00 71.30 8% Zn/Chrome 250 89.06 60.20 300 100.0082.93 350 100.00 72.72 Chrome 250 56.69 46.72 300 98.93 83.35 350 100.0072.31 5% In/Chrome 250 98.64 81.36 300 100.00 68.23 350 100.00 52.65 6%Zn/Chrome 250 96.51 85.17 300 100.00 69.18 350 100.00 72.19 Zn/Chrome250 95.64 82.78 300 100.00 71.90 350 100.00 70.47 Zn/Chrome 250 91.7681.19 300 100.00 77.74 350 100.00 70.90 Chrome 250 93.60 80.31 300100.00 48.69 350 100.00 38.38 Mo/Chrome 250 93.24 80.25 300 100.00 54.87350 100.00 26.63 Ni/Chrome 250 100.00 84.05 300 100.00 39.93 350 100.0024.93 Nb/Chrome 250 98.94 86.54 300 100.00 56.26 350 100.00 34.94Alumina 250 29.10 38.87 300 73.37 69.75 350 98.37 90.94 0.5% Pt/Alumina250 44.53 48.63 300 87.56 82.09 350 100.00 96.02 Fe/Alumina 250 22.4539.46 300 52.57 71.90 350 80.99 84.47 20% Cr/Alumina 250 43.85 44.17 30097.71 79.58 350 98.56 77.88 50% Cr/Alumina 250 45.73 42.68 300 100.0079.86 350 100.00 72.95 Zn/Cu/Alumina 250 40.73 47.65 300 74.85 71.21 35065.17 64.86 0.5% Pd/Carbon 250 100.00 99.19 300 100.00 98.38 350 100.0085.78 activated carbon 175 46.06 99.13 200 98.46 97.32 300 99.07 97.02

All of the catalysts tested were found to be effective at converting243db to 1233xf, particularly activated carbon.

EXAMPLE 2

6.07 g of a Indium-doped chromia catalyst was dried over 72 hours undernitrogen (80 ml/min) at 250° C. and 3 barg. This was followed bytwo-stage pre-fluorination of the catalyst. In stage 1, the catalyst wasexposed to nitrogen (80 ml/min) and HF (4 ml/min) at 250° C. and 3 bargup until 4 hours from HF breakthrough, at which time the temperature wasincreased at 25° C./min to 300° C. and held for 16 hours. In stage 2,nitrogen flow was reduced stepwise until it was switched off, and thetemperature was increased at 25° C./min to 380° C. and held for 10hours. The HF flow was stopped and replaced with nitrogen (40 ml/min)and the temperature reduced to 250° C. ready for use.

1233xf was co-fed with HF over the catalyst without an air co-feed(cycle 1) and with an air co-feed (cycle 2) for about 100 hours at 350°C. and 15 barg. Reactor off-gas was analysed by GC. Monitored catalystregeneration was used to measure the average coke levels in the catalystafter use. The results are shown in Table 2 below and illustrated inFIG. 2.

TABLE 2 Experimental conditions and results for two ageing runs CycleCoke Target flows Conversion time levels (ml/min) Loss Loss rate Cycle(hrs) (%) HF 1233 × f Air (%) Hours (5/hr) 1 100.5 5.6 50 5 — 64.8 740.88 2 110 0.35 45 5 5 16.8 106 0.16

Both cycles were conducted at the same temperature and pressure, but theHF flow was reduced in cycle 2 to maintain contact time. Contact timesfor cycles 1 and 2 were 57 seconds and 65 seconds, respectively. As aresult of the reduced HF flow on cycle 2 and the lower than target1233xf flows, which were hard to control and lower than target, theHF:1233xf ratio differed slightly on the two cycles (average 20:1 forcycle 1 and 15:1 for cycle 2). The average 1233xf flow for cycle 1 was3.2 ml/min and 3.5 ml/min for cycle 2.

Without air the majority of catalyst activity was lost after about 80hours. The introduction of air significantly reduced the rate ofcatalyst deactivation (the activity loss after 100 hours was comparablewith just 20 hours without air). Based on this conversion loss rate,cycle 2 would be expected to take 410 hours to reach the same conversionloss as cycle 1. The reduced rate of catalyst deactivation with airco-feed is in accordance with the catalyst coke levels measured.

The reaction selectively was also affected, total impurity levelsapproximately doubled with the co-feed of air compared to without air.Co-feeding air seemed to have a little impact on the 245cb:1234yf ratiothough.

The concentration of air present was higher than desired because the1233xf flow rate was, on average, lower than the targeted 5 ml/min. Thiswas thought to be at least partially responsible for the decreasedselectivity. For this reason, and based on the coke produced in cycle 1,a lower air concentration is believed to be desirable to achievecomparable reduced rates of catalyst deactivation without reducingconversion and 245cb selectivity. It was estimated that lower air flows,for example from about 0.5 ml/min to about 4.5 ml/min, preferably fromabout 1 to about 4 ml/min, such as from about 1.5 to 3.5 ml/min (allbased on an actual 1233xf flow of 5 ml/min) would realise the surprisingbalance of reduced rates of catalyst deactivation combined withconversion selectivity to the desired 245cb product.

The invention is defined by the following claims.

1-57. (canceled)
 58. A process for preparing1,1,1,2,2-pentafluoropropane (245cb), the process comprising: gas phasecatalytic dehydrochlorination of a composition comprising1,1,1-trifluoro-2,3-dichloropropane (243db) to produce an intermediatecomposition comprising 3,3,3-trifluoro-2-chloro-prop-1-ene (CF₃CCl═CH2,1233xf), hydrogen chloride (HCl) and air; and gas phase catalyticfluorination with hydrogen fluoride (HF) of the intermediate compositionto produce a reactor product composition comprising 245cb, HF, HCl andair; wherein the process is carried out with a co-feed of air, whereinthe amount of air co-fed to the process is from 0.1 to 500 mol %, basedon the amount of organics.
 59. A process according to claim 58 whereinthe dehydrochlorination step is carried out in a first reactor and thefluorination step is carried out in a second reactor.
 60. A processaccording to claim 58 wherein the amount of air co-fed to the process isfrom 1 to 200 mol %, based on the amount of organics.
 61. A processaccording to claim 60 wherein the amount of air co-fed to the process isfrom 2 to 100 mol %, based on the amount of organics.
 62. A processaccording to claim 61 wherein the amount of air co-fed to the process isfrom 5 to 100 mol %, based on the amount of organics.
 63. A processaccording to claim 62, wherein the amount of air co-fed to the processis from 10 to 100 mol %, based on the amount of organics.
 64. A processaccording to claim 63 wherein the amount of air co-fed to the process isfrom 15 to 95 mol %, based on the amount of organics.
 65. A processaccording to claim 64 wherein the amount of air co-fed to the process isfrom 20 to 90 mol %, based on the amount of organics.
 66. A processaccording to claim 65 wherein the amount of air co-fed to the process isfrom 25 to 85 mol %, based on the amount of organics.
 67. A processaccording to claim 59 wherein air is co-fed to both first and secondreactors and wherein the amount of air co-fed to the first reactor isless than the amount, on a molar basis, of air co-fed to the secondreactor.
 68. A process according to claim 67 wherein the amount of airco-fed to the first reactor is less than half the amount of air co-fedto the second reactor.
 69. A process according to claim 68 wherein theamount of air co-fed to the first reactor is less than a quarter of theamount of air co-fed to the second reactor.
 70. A process according toclaim 69 wherein the amount of air co-fed to the first reactor is lessthan a tenth of the amount of air co-fed to the second reactor.
 71. Aprocess according to claim 59 wherein the intermediate composition exitsthe first reactor and is fed directly to the second reactor.
 72. Aprocess for preparing 1,1,1,2,2-pentafluoropropane (245cb), the processcomprising: gas phase catalytic dehydrochlorination in a first reactorof a composition comprising 1,1,1-trifluoro-2,3-dichloropropane (243db)to produce an intermediate composition comprising3,3,3-trifluoro-2-chloro-prop-1-ene (CF₃CC1=CH₂, 1233xf) and hydrogenchloride (HCl); and gas phase catalytic fluorination with hydrogenfluoride (HF) in a second reactor of the intermediate composition toproduce a reactor product composition comprising 245cb, HF, HCl and air;wherein the process is carried out with a co-feed of air to the secondreactor, wherein the amount of air co-fed to the second reactor is from5 to 100 mol %, based on the amount of organics.
 73. A process accordingto claim 72 wherein the amount of air co-fed to the second reactor isfrom 10 mol % to 100 mol %, based on the amount of organics.
 74. Aprocess according to claim 73 wherein the amount of air co-fed to thesecond reactor is from 15 to 95 mol %, based on the amount of organics.75. A process according to claim 74 wherein the amount of air co-fed tothe second reactor is from 20 to 90 mol %, based on the amount oforganics.
 76. A process according to claim 75 wherein the amount of airco-fed to the second reactor is from 25 to 85 mol %, based on the amountof organics.
 77. A process according to claim 72 wherein air isadditionally co-fed to the first reactor and the intermediatecomposition further comprises air.
 78. A process according to claim 77wherein the amount of air co-fed to the first reactor is from 0.1 to 100mol %, based on the amount of organics.
 79. A process according to claim78 wherein the amount of air co-fed to the first reactor is from 0.2 to50 mol %, based on the amount of organics.
 80. A process according toclaim 79 wherein the amount of air co-fed to the first reactor is from0.3 to 20 mol %, based on the amount of organics.
 81. A processaccording to claim 80 wherein the amount of air co-fed to the firstreactor is from 0.4 to 10 mol %, based on the amount of organics.
 82. Aprocess according to claim 81 wherein the amount of air co-fed to thefirst reactor is from 0.4 to 5 mol %, based on the amount of organics.83. A process according to claim 72 wherein the intermediate compositionexits the first reactor and is fed directly to the second reactor.
 84. Aprocess according to claim 83 wherein the catalytic dehydrochlorinationof 243db is carried out in the presence of HF and the intermediatecomposition further contains HF.
 85. A process according to claim 84wherein the composition comprising 243db additionally contains HF, witha molar ratio of HF:243db of from 0.5:1 to 40:1.
 86. A process accordingto claim 85 wherein the molar ratio of HF:243db is from 1:1 to 15:1. 87.A process according to claim 59 wherein molar ratio of HF:1233xf in thesecond reactor of from 1:1 to 45:1.
 88. A process according to claim 87wherein the molar ratio of HF:1233xf in the second reactor is from 2:1to 20:1.
 89. A process according to claim 88 wherein the molar ratio ofHF:1233xf in the second reactor is from 3:1 to 15:1.
 90. A processaccording to claim 87 wherein an additional feed of HF is provided tothe second reactor.
 91. A process according claim 59 wherein the air iscompressed prior to being co-fed.
 92. A process according to claim 59wherein the air is dried prior to being co-fed.
 93. A process accordingto claim 59 wherein the reactor product composition is separated into astream comprising 245cb and HF and a stream comprising HCl and air. 94.A process according to claim 93 wherein the stream comprising 245cb andHF is separated into a 245cb-rich stream and a HF-rich stream.
 95. Aprocess according to claim 94 wherein the 245cb-rich stream is subjectedto a scrubbing step in which residual HF is substantially removed fromthe 245cb-rich stream to produce a 245cb-rich stream substantially freefrom HF.
 96. A process according to claim 59 wherein the 245cb isseparated from any further fluorocarbons present to produce asubstantially pure 245cb product.
 97. A process according to claim 59wherein the catalytic dehydrochlorination is carried out at atemperature of from 200 to 450° C. and a pressure of from 0.1 to 30bara.
 98. A process according to claim 97 wherein the catalyticdehydrochlorination is carried out at a temperature of from 250 to 380°C. and a pressure of from 1 to 20 bara.
 99. A process according to claim98 wherein the catalytic dehydrochlorination is carried out at atemperature of from 300 to 350° C. and a pressure of from 5 to 20 bara.100. A process according to claim 59 wherein the catalyticdehydrochlorination is carried out in the presence of a bulk form orsupported catalyst comprising activated carbon, a zero-valent metal, ametal oxide, a metal oxyhalide, a metal halide, or mixtures of theforegoing.
 101. A process according to claim 100 wherein the metal is atransition metal, an alkaline earth metal or aluminum.
 102. A processaccording to claim 100 wherein the catalyst is based on chromia,preferably a zinc/chromic catalyst.
 103. A process according to claim 59wherein the catalytic fluorination is carried out at a temperature offrom 200 to 450° C. and a pressure of from 0.1 to 30 bara.
 104. Aprocess according to claim 103 wherein the catalytic fluorination iscarried out at a temperature of from 250 to 420° C. and a pressure offrom 1 to 20 bara.
 105. A process according to claim 104 wherein thecatalytic fluorination is carried out at a temperature of from 300 to380° C. and a pressure of from 5 to 20 bara.
 106. A process accordingclaim 59 wherein the catalytic fluorination is carried out in thepresence of a bulk form or supported catalyst comprising activatedcarbon, a zero-valent metal, a metal oxide, a metal oxyhalide, a metalhalide, or mixtures of the foregoing.
 107. A process according to claim106 wherein the metal is a transition metal, an alkaline earth metal oraluminium.
 108. A process according to claim 106 wherein the catalyst isbased on chromia.
 109. A process according to claim 10 wherein thecatalyst is based on a zinc/chromia catalyst.
 110. A process accordingto claim 59 wherein the HF in the reactor product composition is atleast partially recycled to the catalytic dehydrochlorination of thecomposition comprising 243db and HF.
 111. A process according to claim94 wherein the HF in the HF-rich stream is recycled to the catalyticdehydrochlorination of the composition comprising 243db and HF.
 112. Aprocess according to claim 111 wherein the HF-rich stream is separatedinto an HF stream and an organic stream, wherein the HF stream isrecycled to the catalytic dehydrochlorination of the compositioncomprising 243db and HF.
 113. A process according to claim 59 whereinthe reactor product composition further contains2,3,3,3-tetrafluoropropene (1234yf).
 114. A process according to claim59 further comprising feeding the 245cb into a dehydrofluorinationreactor to produce a dehydrofluorination product comprising2,3,3,3-tetrafluoropropene (1234yf) and HF.
 115. A process according toclaim 113 wherein the 245cb is catalytically dehydrofluorinated to1234yf in the gas phase.